Chemical process for hydrocracking and hydrorefining of hydrocarbon oils



CHEMICAL PROCESS FOR HYDROCRACKTNG AND HYDROREFlNlNG F HYDROCAREQN @lLS Harold Eeuther, Gibsonia, and Stephen L. Pealre and Bruce K. Schmidt, Pittsburgh, Pa, assignors to Gulf Research & Development Company, Pittsburgh, Pa., a corporation of Delaware No Drawing. Filed Aug. 25, 1964-, Ser. No. 392,026

19 Claims. (Cl. 208-112) Our invention relates to the hydrogen treatment of heavy petroleum hydrocarbons and derivatives or fractions thereof and in particular relates to the removal of sulfur and nitrogen compounds therefrom and to the hydrocracking thereof.

These heavy petroleum hydrocarbons are available commercially in considerable abundance but they are of relatively low value. In many instances such heavy petroleum hydrocarbons are employed as low-grade fuels without further treatment. This, at best, results in a poo-r return to the producer and the return is being further reduced since many jurisdictions have adopted regulations restricting the amount of sulfur containing material that can be present in fuels burned within their boundaries, thereby requiring removal of substantial quantities of sulfur from these hydrocarbons, such as residuals, prior to sale. It has been previously suggested that sulfur be removed from residual fractions and other heavy petroleum hydrocarbons by subjecting the heavy hydrocarbon to treatment with hydrogen in the presence of hydrogenation catalysts. These suggested procedures, however, entail the employment of expensive operating conditions such as, for example, unusually high pressures, i.e., at least 2000 p.s.i. Previous attempts to remove sulfur from residual stocks by hydrogenation at moderate pressures have resulted in extremely rapid deactivation of the catalysts by metal and coke deposition on the catalysts due to the metalliferous and asphaltic contaminants contained in such stocks. Further, the previous attempts have resulted in an unsatisfactorily low degree of desulfurization. Similarly, attempts to enhance the value of these heavy petroleum hydrocarbons by removal of nitrogen through hydrogenation have been confronted with the same problems.

An alternative course which has been proposed as a means of enhancing the value of these heavy petroleum hydrocarbons is to hydrocrack such stocks in order to produce more valuable, lower-boiling fractions. To effect hydrocracking of heavy hydrocarbons, it is necessary to employ a catalyst which is extremely active in the presence of asphaltic and metalliferous contaminants and which has extremely good aging characteristics, i.e., maintains its cracking activity for long periods of time while exposed to large quantities of asphaltic materials. Such characteristics are not necessarily required when hydrocracking lighter stocks such as, for example, a middle distillate fraction. As mentioned above, however, metal impurities in the stocks to be treated cause rapid deterioration of the activity of the catalysts, and the coking of the catalysts due to the presence of asphaltics in the stocks also has a strong deactivating effect. Thus, a commercially acceptable hy-drocracking process for heavy petroleum hydrocarbons requires not only the employment of an extremely active catalyst in the presence of asphaltic and metalliferous contaminants but also the employment of a catalyst and operating conditions which provide good catalyst aging.

It is an object of our invention to provide improved procedure for the hydrogen treatment of heavy petroleum hydrocarbons.

Another object of our invention is to provide improved 3,322,600 Patented May 30, 1967 procedure for the removal or reduction of sulfur and/ or nitrogen compounds from heavy petroleum hydrocarbons by catalytic hydrogen treatment.

A further object of our invention is to provide improved procedure for hydrodesulfurization of heavy petroleum hydrocarbons employing moderate pressures, e.g., below about 2000 p.s.i.

Another object is to provide improved procedure for hydrocracking heavy petroleum hydrocarbons, particularly a procedure which enables high conversion into lower boiling materials.

These and other objects are accomplished by our invention which includes treating a heavy petroleum hydrocarbon with hydrogen in the presence 'of a catalyst comprising essentially a minor amount of a hydrogenation catalyst composited with a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than A. radius and having less than 10 percent of said pore volume in pores larger than 80 A. radius. This activated alumina is obtained by drying and calcining an alumina sol represented by the formula Al(OH) Z wherein Z is the anion of a salt,

and y is less than 1; the alumina sol being prepared by the reaction of a water soluble aluminum salt and another source of aluminum ions in the presence of water at a temperature above F.

The heavy petroleum hydrocarbons treated in accordance with our invention contain sulfur, asphaltic and metallifero-us compounds as contaminants, contain substantial amounts of hydrocarbon components boiling above 300 F., and contain residual materials. As employed herein the term residual materials means the undistilled petroleum fraction containing the highest boiling components of the crude. There are many natural petroleum hydrocarbons of a heavy nature containing residual materials and containing the contaminants mentioned. All such stocks can be treated in accordance with our invention. Also many petroleum hydrocarbons yield residual fractions on distillation under reduced or atmospheric pressures, which fractions contain the contaminants mentioned. Examples of such stocks are reduced and topped crudes. Our invention is applicable to the treatment of all such residual fractions. Further, our invention is particularly useful in the treatment of petroleum fractions boiling in the range above about 700 F. and even above about 1000 F.

Broadly, the operating conditions employed in the practice of our invention include a hydrogen partial pressure from about 300 to about 3000 p.s.i., a space velocity from about 0.1 to about 10.0 volumes of heavy hydrocarbon per volume of catalyst per hour, a hydrogen feed rate from about 2000 to about 30,000 standard cubic feet per barrel of charge stock (s.c.f./b.), a hydrogen consumption rate from about 1 to about 30 or 40 mols of hydrogen per atom of sulfur removed from the charge stock and a temperature from about 700 to about 900 F. The pressures mentioned throughout the specification and in the claims are to be construed as hydrogen partial pressures whether specifically so stated in each instance or not. The temperatures mentioned herein and in the claims are average reactor temperatures since the diflference between inlet and outlet temperatures can differ as much as 30 F. and even more depending upon the charge stock and the severity of other operating conditions.

As indicated, the activated alumina carrier or support employed in our invention is obtained by drying and calcining an alumina sol represented by the formula Al(OH) Z wherein Z is the anion of a salt. The sol is prepared by the reaction of a water soluble aluminum salt and another source of aluminum ions in the presence of Water. Any water soluble aluminum salt can be employed in preparing the sol, such as, for example, an aluminum-nitrate, -halide, acetate or other carboxylic acid salt. Of the carboxylic acid salts, while the acetate is satisfactory, it is expensive and not quite so reactive as some others. While all of the halide salts of aluminum can be employed in preparing the sol, the fluoride cannot be used alone but must be employed together with at least one other halide. Further, the fluoride salt should only be used when the presence of fluorine in the final catalyst is desired and even then the fluoride salt should only be employed in minor amounts. Gen rally, we prefer to employ aluminum chloride as the water soluble aluminum salt is preparing the sol.

The quantity of water employed in the preparation of the sol should be at least the stoichiometric amount sufficient to provide the quantity of OH ions required by the formula Al('OH) Z wherein x+y=3 and y is less than 1, preferably less than 0.4. A quantity of water in excess of this stoichiometric requirement can also be employed without adversely affecting the formation of the sol and, in fact, a stoichiometric excess of water tends to speed the reaction. Any excess water remaining after the reaction has been completed and the sol formed can readily be driven off during the drying step. I

The source of aluminum ions other than the water soluble aluminum salt described above can be aluminum hydroxide, aluminum hydrate, freshly formed aluminum hydrogel or aluminum metal, such as, for example, aluminum scrap. The employment of aluminum scrap provides a comparatively inexpensive source of reasonably pure aluminum which generally does not require extensive cleaning prior to use. A certain degree of caution should be exercised when employing aluminum metal as a starting material since a large amount of hydrogen is evolved in the reaction which is exothermic. Thus, it is possible for the reaction to become extremely violent. Generally we prefer to employ aluminum metal as a starting material.

After combining the starting materials, the reaction mixture is heated to a temperature above about 120 F. and preferably to a temperature in the range from about'160 F. up to the boiling point of the mixture. In some instances the heating can be discontinued, or at least reduced, after the reaction has commenced due to the exothermic nature of the reaction; provided that the reaction mixture is maintained at a temperature within the ranges mentioned above.

In the course of preparing the alumina sol once the reactants have been brought together and the mixture has been heated .to the reaction temperature it is essential that the reaction mixture be agitated throughout the course of the reaction. After heating the mixture to the desired temperature, a vigorous reaction ensues with the liberation of hydrogen. The length of time required for the reaction can vary widely depending upon the selection of starting materials and the reaction temperature with variations from several hours up to several days not being uncommon. When employing a hydrogel as a starting material, the reaction is extremely rapid and is usually complete within a few minutes. Upon completion of the reaction a residue of unreacted aluminum metal is found in the reaction vessel.

The alumina sol thus produced can then be dried by any of the methods well-known in the art, such as heating to drive off the water, to yield a transparent and glassy aluminum hydroxide which X-ray diffraction analysis shows to be a less well crystallized fine grain boehmite.

The starting materials described above when employed in the proportion required to produce the sol of the formula shown previously will generally provide a reaction mixture having a pH in the range from about 3 to about 4 or 4.5. It is, however, essential to this preparation that the pH of the reaction mixture be maintained at 4.5 or below since a pH greater than 4.5 causes the formation of an alumina hydrogel and precludes the formation of the desired sol.

It will also be noted that in the formation of the sol and the subsequent formation of the alumina the presence of a mercury activator is not required. The absence of the mercury activator not only contributes to the formation of the particular alumina required in the process of our invention but eliminates the necessity of subsequently removing the mercury before forming the final catalyst.

After drying, the alumina is calcined to obtain an activated alumina which constitutes the catalyst support employed in our invention. Any conventional method heretofore used for calcing a dried alumina may be employed. However, a temperature above about 1600 F. should not be used since such elevated temperatures cause deactivation of the activated alumina. A temperature of between about 800 F. and 1200 F. ordinarily is satisfactory. A calcining time of between about 2 and 24 hours ordinarily will be satisfactory. In most cases the shorter time periods will be used with the higher temperautres and the longer periods with the lower temperatures. The final product is opaque and hard. It has a unique pore structure and is A1 0 which still contains a small amount of waterusually less than about 3 percent.

This unique pore structure of the alumina used in our process can be characterized as including a substantial, usually a predominant, portion of the total pores and consisting of pores less than 300 A. radius having the particular pore size distribution described previously. For example, a similar alumina has been found to be composed almost entirely of pores less than 300 A. radius, which pores comprise more than percent of the total pore volume.

In the practice of our invention we employ the activated alumina described above as the catalyst support rather than a commercial alumina-containing silica support inasmuch as it provides the desired activity but does not possess the extremely high initial activity of commercial silica-alumina supports and, therefore, does not cause extremely high initial cracking which in turn causes rapid coking of the catalyst.

Surprisingly, we have found that when this activated alumina is composited with suitable metals having catalytic activity, a catalyst composition is obtained which yields unexpectedly superior results in the hydrocracking and hydrodesulfurization of heavy petroleum hydrocarbons. Such activity is particularly unexpected in light of the fact that, when the alumina described above is substituted for commercially available aluminas of a different type in certain other hydrocarbon treatment processes, e.g., hydrocracking of distillate stocks, the results achieved are no better than and, in some instances, even inferior to the results obtained employing the commercial supports.

The activated alumina carrier described above is composited with a metalliferous hydrogenating component. Any of the conventional procedures for preparation of such a two-component catalyst may be used. Ordinarily, we prefer to impregnate the activated alumina carrier with an aqueous solution of a salt of the metalliferous hydrogenating catalyst and then dry and calcine to obtain the finished hydrogenation catalyst. Any hydrogenating component such as Group VIII or Group VI metal oxides or sulfides such as molybdenum or tungsten oxides and sulfides or nickel or cobalt metals, oxides or sulfides may be used. It is frequently desirable to employ mixtures of these catalysts such as cobalt-molybdenum, nickel-cobaltmolybdenum, nickel-tungsten, etc., their oxides or sulfides.

A particularly desirable catalyst is a mixture of oxides of nickel, cobalt and molybdenum such as described in US. Patent 2,880,171, Mar. 31, 1959, Flinn et al.

In one aspect our invention relates to the desulfurization of heavy petroleum hydrocarbons. When operating in accordance with this aspect of our invention, the contacting of the charge stock with hydrogen is carried out under conditions of temperature and space velocity which avoids substantial or extensive cracking of carbon-tocarbon bonds. We have found that by operating in accordance with this procedure moderate pressures of between about 300 and 2000 p.s.i. can be employed yet nevertheless long onstream periods or high throughputs may be used to give extensive desulfurization. This procedure not only results in improved removal of sulfur compounds but also in an improved removal of nitrogencontaining impurities. Other advantages such as lower coke deposits and lower carbon residue are also obtained.

Generally, temperatures within the range from about 700 to about 875 F. and preferably between about 750 and 850 F. are employed. By employing temperatures within these ranges is meant to commence operation at the lower end of the temperature range and gradually increase the temperature during the course of operation in order to maintain the rate of desulfurization or the sulfur content of the product constant until a terminal temperature is reached, at which time the reactor is shut down and the catalyst regenerated. Thus, for example, the hydrodesulfurization process can be commenced at a temperature within the range from amout 700 to 800 F. and the temperature increased incrementally until a temperature in the range of about 775 to 875 F. is attained, preferably the process can be commenced at a temperature of about 750 to 775 F. and the temperature increased gradually until a temperature of about 800 to 850 F. is attained at the end of the run.

Also when practicing this aspect of our invention a hydrogen partial pressure between about 300 and 2000 p.s.i. and preferably between 1000 and 1500 p.s.i. is used, a space velocity between about 0.10 and 10.0 and preferably between 0.2 and 3.0 can be used, a hydrogen feed rate from about 2.000 to about 20,000 s.c.f./ b. and preferably from about 4000 to about 10,000 s.c.f./b. is employed, and a hydrogen consumption of about 1 to about 5 mols of hydrogen per atom of sulfur removed from the charge stock and preferably from about 2 to about 5 mols of hydrogen per atom of sulfur is employed. While hydrodesulfurization causes the formation of lower boiling materials due to the rupture of the sulfur bonds in the hydrocarbon molecule, this type of rupturing does not cause deposition of coke on the catalyst and is desirable. Since the process of this aspect of our invention results in extensive rupturing of sulfur bonds, there will be lower boiling hydrocarbons formed as a by-product and these hydrocarbons have desirable properties as compared with the residual material being treated. Since the objective is the utilization of as long an onstream period as possible, it is best to use conditions which avoid carbon-tocarbon cracking inasmuch as this results in deposition of coke on the catalyst and this in turn increases the rate of deactivation of the catalyst. Therefore, this cracking should be maintained below about 20 percent formation of volatile material over and above that formed by rup turing of sulfur bonds. The space velocity and temperature can be regulated to give the desired mild conditions which avoid extensive carbon-to-carbon cracking. Thus, with a given catalyst the higher the temperature and the lower the space velocity the higher will be the amount of carbon-to-carbon cracking whereas the reverse conditions utilizing a lower temperature in the range given and a higher space velocity will reduce the carbonto-carbon cracking. If this carbon-to-carbon cracking is kept below the above mentioned maximum value, the reaction may be continued for relatively long periods of time on the order of 3 to 75 barrels of residual feed stock per pound of catalyst. Eventually the catalyst will require regeneration and this is accomplished in the usual fashion by terminating the onstream reaction and burning the carbonaceous material from the catalyst by combustion regeneration. The regenerated catalyst then may be reused in the process.

Of the catalysts mentioned above it is preferred to utilize a catalyst of the cobalt molybdate type in practicing this aspect of our invention since such catalyst has a high desulfurization activity and a low activity for carbon-tocarbon splitting. As mentioned above, a particularly desirable catalyst of this type is a mixture of the oxides of nickel, cobalt and molybdenum as described in US. Patent 2,880,171. Other nickel-cobalt-molybdneum catalysts in which the total metal content is less than about 20 to 25 percent by weight of the catalyst and in which the atomic ratio of Group VIII to Group VI metals is greater than 1.0 can also be employed with equally satisfactory results.

Another aspect of our invention relates to the hydrocracking of heavy petroleum hydrocarbons. When operating in accordance with this aspect of our invention the heavy hydrocarbon is contacted with hydrogen under hydrocracking conditions of temperature and pressure in the presence of a two-component catalyst. The hydrogenating component of this catalyst comprises a Group VI metal together with a Group VIII metal or their oxides or sulfides. Thus, for example, mixtures of these components such as nickel cobalt-molybdenum, cobaltmoylbdenum, nickel molybdenum, nickel tungsten, etc., and their oxides and sulfides can be employed. The support for the hydrogenating component is the activated alumina described above. The activated alumina is composited with the hydrogenating component in accordance with any of the conventional procedures for impregnation of porous carriers with multi-component catalysts. Ordinarily, we prefer to impregnate the activated alumina carrier with an aqueous solution of a salt of the Group VI metal such as molybdenum followed by drying and calcining and then to impregnate with an aqueous salt of a Group VIII metal such as nickel or cobalt followed by a second drying and calcining. If desired the oxides of the metal components can be reduced or partially reduced by treatment with hydrogen prior to employment in our process. In the event a sulfide is to be present, the catalyst can be treated with hydrogen sulfide to form the metal sulfides. This is advantageously carried out by treating with a mixture of hydrogen and hydrogen sulfide at a temperature between about 450 and 950 F. Also the.

metal components can be reduced and/ or sulfided by contacting with the feed stock. Alternatively, the catalyst may be formed by precipitating the sulfides of the metals in aqueous impregnating solutions as by treatment with hydrogen sulfide.

In accordance with this aspect of our invention the heavy hydrocarbon to be hydrocracked is contacted with the above described catalyst at a temperature between about 750 and 900 F. and preferably between 780 and 875 F. By employing temperatures within these ranges is meant to commence operation at the lower end of the temperature range and gradually increase the temperature during the course of operation until a terminal temperature is reached, at which time the reactor is shut down and the catalyst regenerated. Thus, for example, the hydrocracking process can be commenced at a temperature within the range from about 750 to about 850 F. and the temperature increased incrementally until the temperature range of about 810 to about 900 F. is attained, at which point the run is terminated. Preferably, the process can be commenced at a temperature within the range from about 780 to about 810 F. and then increased incrementally until a temperature in the range from about 840 to about 875 F. is attained. The tem perature is increased during the course of the process from the lower starting range at a rate sufiicient to maintain the volume of distillate yield at a predetermined satisfactory level and this gradual increase of temperature is continued until such time as the upper limit of the temperature range has been achieved.

When operating in accordance with this aspect of our invention a hydrogen partial pressure between about 1500 and 3000 p.s.i., preferably between about 2000 and 2500 p.s.i., a space velocity from about 0.1 to 5.0, preferably from about 0.2 to 2.0, a hydrogen feed rate from about 5000 to about 30,000 s.c.f./b., preferably from about 7000 to about 20,000 s.c.f./b, and a hydrogen consumption rate from about 6 to about 40 mols of hydrogen per atom of sulfur removed from the charge stock, preferably from 6 to 20 mols of hydrogen per atom of sulfur, can be employed.

The characteristics of the products obtained from the practice of this aspect of our invention will depend upon the feed stock, particularly the boiling point of the feed stock, and the reaction conditions employed. Thus, it is possible to produce a furnace oil product which is usable directly from the processing unit. Furthermore, this aspect of our invention can also be employed to produce a low octane gasoline directly from residual stocks. In the practice of this aspect of our invention it is also possible to recycle to the reactor all portions of the product boiling above a particular range, such as, for example, the gas oil range (1050 F.), thereby increasing the net yield of usable products per volume of heavy hydrocarbons charged.

As mentioned previously, the pore size distribution of the activated alumina employed in our invention in the areas of both hydrocracking and hydrodesulfurization of heavy petroleum hydrocarbons yields unexpectedly superior results. It is theorized that these unexpectedly superior results are due principally to the extremely small pore size of the support. As is well known in the art, heavy petroleum hydrocarbons, such as residual stocks, for example, contain large quantities of asphaltic and metalliferous materials and that the presence of such contaminants adversely affects catalyst life; the asphaltics by depositing coke on the catalyst and the metalliferous materials by depositing metals on the catalyst surface thereby poisoning the catalyst. As is also well known, the asphaltic compounds are polyaromatic molecules of comparatively large size and the metals present in petroleum stocks are normally contained in the form of extremely large molecules. We believe that the mechanics of the hydrocracking and hydrodesulfurization aspects of our invention are such that, since a substantial portion of the surface area of the catalyst is within the extremely small pores of the catalyst, a great number of-the extremely large molecules present in the stock are prevented from entering the small pores where they might be adsorbed on the catalyst surface and react. Thus, the particular type of activated alumina described above when employed in our invention permits only a limited amount of the large molecular asphaltic and metalliferous materials to react while providing an active catalyst of high surface area suitable for adsorbing the comparatively smaller hydrocarbon molecules and permitting them to react. It is believed, therefore, that minimizing the reaction of the asphaltic and metalliferous materials in this manner also minimizes the coke and metal deposition on the surface of the catalyst without any reduction in the desired reactions. A comparison of the products obtained in accordance with our invention with those obtained by known processes tends to substantiate this theory, inasmuch as the products of our process have a higher metals content and contain more higher boiling asphaltics. We believe, therefore, that this theory accounts for the unique co-action of the particular activated alumina described above with heavy petroleum hydrocarbons in accordance with our invention.

In order to illustrate our invention in greater detail, reference is made to the following examples.

8 Example I A gram quantity of aluminum chloride (AlCl 6H OMallinckrodt, reagent grade) was dissolved in 1000 ml. of distilled water. Into each of two 6-liter Erlenmeyer fiasks were placed 500 grams of aluminum metal granules (total=1000 grams of aluminum) together with 3 liters of distilled water. To each of the two flasks were then added 100 ml. of the aqueous aluminum chloride solution and the mixtures were heated to F. while continuously stirring the mixture. A vigorous evolution of hydrogen was obtained indicating that the reaction had commenced. Heating of the reaction mixture was then discontinued and the temperature was maintained at 160 F. by the heat of the reaction. As the reaction in the flasks slowed down somewhat, portions of the aqueous aluminum chloride solution were added to each of the flasks to promote the reaction until 500 ml. of the aluminum chloride solution had been added to each of the Erlenmeyer flasks. After the reaction had terminated, there remained an aluminum sol of the approximate formula Al(OH) Cl and about 420 grams of unreacted aluminum metal. The sol was then evaporated to dryness on a hot plate, dried in an oven at 250 F. and calcined at 900 F. for 16 hours. The activated alumina thus obtained was transparent and glassy. This material was ground to 10-20 mesh size particles, which were impregnated with metals to produce the finished catalyst.

To effect the impregnation an ammonium molybdate solution was prepared by dissolving 22.45 grams of ammonium paramolybdate [(NH4)6MO']O24'4H20] in distilled water and 10 ml. liters of ammonia (28% NI-I and then diluting to 99 ml. with distilled water. The final solution weighed 116.5 grams and contained 10.5 percent molybdenum. This solution was then added with stirring to an evaporating dish which contained 131.3 grams of the alumina support. The solution completely wet the alumina support (incipient wetncss0.753 ml. of solution per gram) and left no excess solution in the dish. The wet material was dried at about 250 F. for 24 hours.

A nickel nitratecobalt nitrate solution was prepared by combining 4.86 grams of 20 percent nickel oxide stock solution (NiNO -6H O and distilled water) and 7.55 grams of cobalt nitrate [Co(NO 6H O] dissolved in distilled water and then diluting this solution with distilled Water to 94.5 ml. The final solution weighed 102.3 grams and contained 0.74 percent nickel and 1.49 percent cobalt. This solution was then added with stirring to an evaporating dish containing 178.5 grams of dried material from the molybdenum impregnation above. This solution completely wet the molybdenum impregnated material (incipient wetness0.53 ml. of solution per gram) and left no excess solution in the dish. The wet material was dried at about 250 F. for 24 hours and calcined in air in an electric mufile furnace at 1000 F. for about 16 hours. The final metal content of the catalyst was about 0.5 percent nickel, 1.0 percent cobalt and 8.0 percent moylbdenum.

To illustrate the process of our inventon, a Kuwait vacuum residue (18 percent by volume of crude) containing sulfur, asphaltic and metalliferous contaminants was subjected to hydrogen treatment in accordance with the process of our invention employing the nominal conditions of 1000 p.s.i.g., 750 F., 0.5 LHSV and 10,000 s.c.f. of hydrogen per barrel of feed. In order to afford a basis for comparison of our invention with previously suggested procedures this hydrogen treatment was carried out employing three different catalysts including the catalyst required in the process of our invention and described immediately above. Of the remaining two catalysts, one was a commercial hydrogenation catalyst which contained 0.3 percent nickel, 1.8 percent cobalt and 10 percent molybdenum deposited on a low density alumina. The

other catalyst was a commercial hydrogenation catalyst which contained 0.5 percent nickel, 1 percent cobalt and 8 percent molybdneum deposited upon a commercial alumina which is widely used in the United States as a carrier for hydrogenation catalysts. The physical characteristics of each of the three catalysts employed are given in Table I below and refer to the calcined support prior to employment in the reaction zone. The pore size distributions shown in Table I were determined by the technique of nitrogen adsorption and desorption isotherms described in the article by Ballou and Doolen in Analytical Chemistry, vol. 32, p. 532, April 1960.

TABLE I Evaporated Commercial Low S01 A1203 A1203 DCllSlLY Surface Area, M /g. 261 172 105 Pore Volume, cc./g 0.38 0.471 0. 446 Avg. Pore Radius (2V/A), 20.0 55 O 45. 7 Tap Density, g./cc 0. 80 0. 707 O. 532 Pore Size Distribution,

Percent of Pore Volume in Range:

200-300 A. Rad 0. 9 0. 9 3. 2 3. 8. 9 10.0 2. 2 17. 2 7. 2 4. 5 21. l 19. 0 4. 6 17. 0 14. 3 11.2 8.2 17. 3 27. 2 13. 8 l3. 2 31. 6 12. 8 12. 7 13. 4 0. 0 2. 7 1.6 0.0 0. 0 0.0 0.0 0.0 3. 9 9. 8 13. 8 0. 1 27.0 21.0

In Table II, column 1 shows the inspection data of the Kuwait vacuum residue employed as charge stock, columns 2 and 3 show the results covering two different periods of time of the run embodying the process of our invention, while columns 4 and 5 show the results obtained in the runs employing the low density alumina and the commercial alumina, respectively.

TAB LE II 10 (column 3). Thus, the percent desulfurization obtained during the period of 24 to 104 hours in accordance with our invention, as shown in column 3, is 76.5 percent, while the runs of columns 4 and 5 for the period of 24-96 hours show a percent desulfurization of only 69.4

1 and 66.2, respectively.

The true advantageous results unexpectedly provided by the process of our invention are more clearly evidenced by the comparison of the percent desulfurization obtained at 30, 60 and 90 hours. It will be noticed that Catalyst Support Evaporated Sol A1203 Column N 0 1 2 Operating Conditions:

Yield, percent by vol. of Charge:

Gasoline (Ci-400 F.) Light Gas Oil (400670 F.) Heavy Gas Oil (6701000 F.) 070 F. Residue- 1000 F. Residue.

Catalyst Deposits: Carbon, percent by Hydrogen Consumption:

S.c.f./bbl Mols Iii/Atom S Removed.

Desulfurization, percent by wt 80.7 Total Liquid Product Inspections Gravity, API 17.0 Viscosity, SUV, 0

100 F.- 2,885 130 F- 962 210 F 22, 130 Sulfur, percent by w 5. 45 1.05 Nitrogen, percent by wt 0.43 0.31 Carbon Residue, Conradson, percent by wt. 23.11 10.23 Insoluble in n-Pentane, percent by wt 15. 14 Vanadium, p.p.1n 102 22. 2 Nickel, ppm 32 12.2 Percent Desulfurization at:

30 hours 79. 8 00 hours. 78. 0 90 hours 70, 2 Deactivation Rate:

Percent Sulfur in Product/100 hrs 0, 33 Percent Desulturization/IOO hrs 6. 1

the percent desulfurization obtained with the process of columns 2 and 3 decreased only 3.6 percent during the period of 30 to 90* hours, while the percent desulfurization decreased about 10 percent during similar periods when employing other processes as shown in columns 4 and 5. A more dramatic illustration of this point is the fact that the percent desulfurization obtained after 90 hours operation in accordance with our invention is somewhat superior to the percent desulfurizati-on obtained after only hours operation but employing dif- 30 ferent types of catalyst supports as shown in columns 4 and 5. From a plot of these and other data it can also be calculated that the rate of deactivation of the catalyst in accordance with the present invention is substantially less than that obtained with commercial cata- 39 lysts (0.33 as opposed to 0.85 and 1.0 percent sulfur for 100 hours). Thus, it can be seen that the life of the catalyst employed in our invention is much greater than the life achieved with commercial catalysts.

Commercial Commercial Low 1) ensity A1203 B After Soxhlet Extraction with benzene.

It will also be noted that the process of our invention provides a greater increase in product gravity in degrees API (17.0 and 16.5) than is provided by the processes employing commercial aluminas (14.6 and 14.7). A comparison of the n-pentane insolubles along with the metals content of the products clearly substantiates the theory of operation of the evaporated sol alumina catalyst support in accordance with our inventoin. Thus, the n-pentane insolubles content of the product of our process shown in column 3 is substanitally greater than that present in the processes employing commercial alumina (columns 4 and 5). A similar comparison of the metals content of the products shows that a far greater quantity of metalliferous contaminants are passed through the catalyst bed during comparable periods of operation when employing the process of our invention (column 3) as opposed to the processes employing commercial aluminas (columns 4 and 5). As stated previously, we believe that the particular alumina employed in accordance with our invention permits a greater portion of the asphaltie and metalliferous contaminants to pass through the reactor bed unconverted, inasmuch as these large molecular contaminants are precluded from entering the smaller size ports of the catalyst. Again, the smaller coke deposition on the catalyst employed in our invention as opposed to the quantity of coke deposition on the commercial alumina catalysts can also be seen by a comparison of the data in Table II.

In a comparison of the data of Table I, it will be noted that the commercial catalysts (columns 2 and 3) which do not possess the pore size distribution required by our invention (less than 5 percent of the pore volume in pores larger than 100 A. and less than percent of the pore volume in pores larger than 80 A.) do not produce the superior results achieved by our invention requiring the use of the catalyst employed in the run illustrated in columns 2 and 3 of Table II.

Example 11 In this example the same residual stock used in Example I is employed as the charge stock to a hydrodesulfurization process. The catalyst employed in this example is also the same nickel-cobalt-molybdenum catalyst supported on alumina derived from an evaporated sol that was used in Example 1. After completion of normal start-up procedures, which include introducing the charge I stock to the reactor at a low temperature, usually well below operating temperatures, and then gradually increasing the temperature over a period of time, usually about 6 to 8 hours, until a normal operating temperature is reached, the run of this example is commenced at an initial operating temperature of 750 F. Throughout the duration of the run the other operating conditions are maintained constant at a pressure of 1000 p.s.i.g., a space velocity of 0.64 LHSV and a hydrogen recycle of 5000 s.c.f./b. The sulfur content of the total liquid product obtained initially is 1.5 percent by weight. During the course of this run the temperature is increased at an average rate of approximately 0.75 F./ day in order to maintain the level of the sulfur content in the total liquid product at 1.5 percent or below. This operation is continued until an operating temperature of 840 F. is achieved, at which time the operation is discontinued. During the course of this run the hydrogen consumption is about 4.0 mols of hydrogen per atom of sulfur removed fromthe charge stock.

Thus, when operating in accordance with the particular embodiment of our invention described immediately above, it will be seen that a process for desulfurizing a heavy petroleum hydrocarbon while operating at a low pressure is provided in accordance with this aspect of our invention. Further, the gradual increase in the operating temperature, in this example 0.75 -F /day, provides a product with a constant low sulfur content, i.e., 1.5 percent 12 or less, while also providing an unexpectedly long catalyst life.

Example III In this example the same residual stock used in Example I was subjected to hydrocracking. In order to form a basis for comparison of this aspect of our invention with prior processes, several separate runs were conducted employing catalysts supported on the evaporated sol alumina required by our invention and employed in Example I, as well as several catalysts supported on commercial aluminas widely used both in the United States and in Europe as carriers for hydrogenation catalysts. Each of the catalysts employed in the runs shown in Table III contained 0.5 percent nickel, 1 percent cobalt and 8 per-cent molybdenum deposited on the alumina supports. The results of the run employing the alumina in accordance with our invention are shown in column 1, while the results of the runs employing the commercial aluminas as supports are shown in columns 2, 3 and 4. In all of the runs the operating conditions were maintained at 790 F., 0.5 LHSV, 2000 p.s.i.g. and 10,000 s.c.f. H b. The data in Table III are the results covering the onstream interval of 8 to 40 hours for all of the runs.

TAB LE III Run No.

Alumina Used {or Catalyst Support;

Evan Commercial Charge S Aluminas Inspections of Liquid Product:

Gravity, API 5. 5 24.0 23. 5 24. 3 25. 3 Viscosity at 130 F., SUS 22, 130 97. 7 98. 4 81). 5 Sulfur, Percent 5. 5 0.25 .25 -4 .22 Nitrogen, Percent 0. 43 0.17 .165 .13 .100 Carbon Residue, Percent 23. 1 3. 22 3. 3. 50 2. 74 Insol. in u-Pentane, Percent 15. l 1. 38 2. 22 1. 50 0. 01 Vanadium, p.p.m 102 0. 7 0. 4 0. 5 0.2 Nickel, p.p.1n 32 0.4 0.2 0. 7 0. 2 Conversion, vol. Percent 1,000 F 62 66 05 67 Carbon on Catalyst, Percent by w 10. 4 16. G 14. 7 15. 6 H2 Consumption:

s.e.i 1, 352 1, 272 l, 034 1, 540 Mols Hz/AtOIIl S Removed. G. 5. 73 7. 34 (1. 92 Pore Size Distribution,* Percent of Pore Vol. in Range:

ZOO-300 A. Radius 0.9 1. 7 2. 5 3. 5 3. 0 10. 4 0. 5 7. 9 2. 2 14. 7 0. 7 4. 8 4. 5 20. 9 10. 7 l0. 8 4. 0 10. 7 14. 7 9. 0 11.2 11. 4 22. 3 14. 7 27. 2 10. 2 l7. 0 l0. 4 31. 0 12. 2 10. 1 21. 3 15. 0 7. 8 0. 0 8. 1 0. 0 0. 0 0. 0 0. 0 3. 9 12. l 12. 0 11. 4 6. 1 26. 8 18. 7 1G. 2

a At 210 F. *Pore Size Distribution determined by nitrogen absorption and desorption isotherms.

It will be noticed that the product obtained in the run employing the alumina of unique pore structure in accordance with our invention shown in column 1 is in every way comparable to, if not superior to, the products obtained when using commercial alumina supports as in columns 2, 3 and 4, with the exception of metal contents. Thus, for example, the API gravity and the viscosity of the products are comparable and the degree of conversion of the processes are comparable. The efiicacy of our invention in reducing catalyst coking is indicated by a comparison of the carbon on the catalyst shown in column 1 13 as opposed to the amount of carbon on catalyst indicated in columns 2, 3 and 4. It can be seen, therefore, that the practice of our invention substantially reduces the amount of coke deposited on the catalyst, thereby providing better catalyst aging and enhancing catalyst life to an un' expected extent.

Example IV In this example the same residual stock used in Example I is employed as the charge stock to a hydrocracking process. The catalyst employed in this example is also the same nickel-cobalt-molybdenum catalyst supported on alumina derived from an evaporated sol that was used in Examples I and 111. After completion of normal startup procedures, which include introducing the charge stock to the reactor at a low temperature, usually well below operating temperatures, and then gradually increasing the temperature over a period of time, usually about 6 to 8 hours, until a normal operating temperature is reached, the run of this example is commenced at an initial operating temperature of 790 F. Throughout the duration of the run the other operating conditions are maintained constant at a pressure of 2000 p.s.i.g., a space velocity of 0.5 LHSV and a hydrogen recycle rate of 10,000 s.c.f./b. The initial conversion of the charge stock to distillate boiling at less than 1000 F. when employing the initial conditions is 65 percent by volume. During the course of this run the temperature is increased at an average rate of approximately 0.75 F./day in order to maintain the level of conversion at 65 percent by volume of the feed to distillate boiling at less than 1000 F. This operation is continued until an operating temperature of 880 F. is achieved, at which time the operation is discontinued. During the course of this run the hydrogen consumption is about 7.5 mols of hydrogen per atom of sulfur removed from the charge stock. Further the product obtained has a substantially higher API gravity (about than the charge stock and a high metals content. The low weight of coke deposit on the catalyst also provides good catalyst aging and an unusually long catalyst life with such feed stock.

Thus, when operating in accordance with the particular embodiment of our invention described immediately above, it will be seen that a process for hydrocracking a heavy petroleum hydrocarbon while operating at a low pressure is provided in accordance with this aspect of our invention. Further, the gradual increase in the operating temperature, in this example 0.75 F./ day, provides a constant level of conversion, while also providing an unexpectedly long catalyst life.

Example V In this example two diiferent hydrocarcking runs were made employing a straight run furnace oil obtained from a Coastal B-l crude. A 6 percent nickel, 19 percent tung sten, 2 percent fluorine supported catalyst was employed in both runs. In one run the support was a commercial alumina (Filtrol Grade 86) While in the other run the support was an alumina derived from an evaporated sol of this invention. The operating conditions employed in both runs included a temperature of 800 F., a pressure of 500 p.s.i.g., a liquid hourly space velocity of 1.0 and a hydrogen feed rate of 10,000 s.c.f./b. The inspections of the furnace oil charge included Gravity, API 29.6

1.4 The conversion in each of these runs Was determined as the percent by volume of product boiling at less than 400 F. The conversions obtained were as follows:

Commercial Al O 66%;

Evap. sol Al O 43%.

From the above data it can readily be seen that when the alumina catalyst support of our invention is employed in a process other than that of our invention, i.e., hydrocracking furnace oil which is not a heavy hydrocarbon containing sulfur, asphaltic and metalliferous compounds, the unique co-action and unexpectedly superior results demonstrated in the previous example are absent.

We claim:

1. A process for hydrogen treatment of heavy petroleum hydrocarbons containing sulfur, asph-altic and metalliferous compounds as contaminants, containing substantial amounts of hydrocarbon components boiling above 300 F. and containing residual materials, which process comprises contacting the hydrocarbons with hydrogen at a hydrogen partial pressure from about 300 to about 3000 p.s.i., a space velocity from about 0.1 to about 10.0 volumes of heavy hydrocarbon per volume of catalyst per hour, a hydrogen consumption rate from about 1 to about 40 mols of hydrogen per gram atom of sulfur removed from the heavy hydrocarbons and at a temperature from about 700 to about 900 F. in the presence of a catalyst comprising essentially a minor amount of a hydrogenating catalyst composited With a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than 100 A. radius and having less than 10 percent of said pore volume in pores larger than A. radius prepared by drying and calcining an aluminum sol having the formula Al(OH) Z wherein Z is the anion of a water soluble aluminum salt, x+y=3 and y is less than 1, said alumina sol being prepared by the reaction of said salt and another source of aluminum ions in the presence of Water at a temperature above about 120 F. and at a pH below about 4.5.

2. The process of claim 1 wherein the hydrogenating catalyst is selected from the group consisting of Group VI metals, Group VIII metals, their oxides and sulfides.

3. The process of claim 1 wherein the hydrogenation catalyst consists essentially of the oxides of nickel, cobalt and molybdenum.

4. The process of claim 1 wherein the hydrocarbon components boil above 1000 F.

5. The process of claim 1 wherein y is less than 0.4.

6. A process for hydrodesulfurizing heavy petroleum hydrocarbons containing asphaltic and metalliferous compounds as contaminants, containing substantial amounts of hydrocarbon components boiling above 300 F., containing residual materials, and containing harmful amounts of sulfur compounds, which process comprises contacting the hydrocarbons with hydrogen at a temperature from about 700 to about 875 F., a hydrogen partial pressure from about 300 to about 2000 p.s.i., a space velocity from about 0.1 to about 10.0 volumes of heavy hydrocarbon per volume of catalyst per hour, a hydrogen feed rate from about 2000 to about 20,000 s.c.f./b., and a hydrogen consumption rate from about 1 to about 5 mols of hydrogen per gram atom of sulfur removed from the heavy hydrocarbons, the contacting rbeing conducted in the presence of a catalyst comprising essentially a minor amount of hydrogenating catalyst composited with a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than A. radius and having less than 10 percent of said pore volume in pores larger than 80 A. radius prepared by drying and calcining an alumina sol having the formula Al(OH) Z wherein Z is the anion of a Water soluble aluminum salt, x+y=3 and y is less than 1, said alumina sol being prepared by the reaction of said salt and another source of aluminum ions in the presence 15 of water at a temperature above about 120 F. and a pH below about 4.5.

7. The process of claim 6 wherein the contacting temperature is in the range from about 750 to about 850 F., the pressure is from about 1000 to about 1500 p.s.i., the space velocity is from about 0.2 to about 3.0 volumes of heavy hydrocarbon per volume of catalyst per hour, the hydrogen feed rate is from about 4000 to about 10,000 s.c.f./b. and the hydrogen consumption rate is "from about 2 to about 5 mols of hydrogen per gram atom of sulfur removed.

8. The process of claim 6 wherein the hydrogenating catalyst is selected from the group consisting of Group VI metals, Group VIII metals, their oxides and sulfides.

9. The process of claim 6 wherein the hydrogenation catalyst consists essentially of the oxides of nickel, cobalt and molybdenum.

10. The process of claim 6 wherein y is less than about 0.4. v

11. A process for hydrocracking heavy petroleum hydrocarbons containing sulfur, asphaltic and metalliferous compounds as contaminants, containing substantial amounts of hydrocarbon componentsboiling above 300 F., containing residual materials, which process comprises contacting the hydrocarbons with hydrogen at a temperature from about 750 F. to about 900 F., a hydrogen partial pressure from about 1500 to about 3000 p.s.i., a space velocity from about 0.1 to about 5.0 volumes of heavy hydrocarbon per volume of catalyst per hour, a hydrogen feed rate from about 5000 to about 30,000 s.c.f./'b. and a hydrogen consumption rate from about 6 to about 40 mols of hydrogen per gram atom of sulfur removed from the charge stock, the contacting being conducted in the presence of a catalyst comprising essentially a minor amount of hydrogenating catalyst composited with a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than 100 A. radius and having less than 10 percent of said pore volume in pores larger than 80 A. radius prepared by drying and calcining an alumina sol having the formula Al(OH) Z wherein Z is the anion of a water soluble aluminum salt, x-t-y=3 and y is less than 1, said alumina sol being prepared by the reaction of said salt and another source of aluminum ions in the presence of water at a temperature above about 120 F. and a pH below about 4.5.

12. The process of claim 11 wherein the contacting temperature is from about 780 to about 875 F., the hydrogen partial pressure is from about 2000 to about 2500 p.s.i., the space velocity is from about 0.2 to about 2.0 volumes of heavy hydrocarbon per volume of catalyst per hour, the hydrogen feed rate is from about 7000 to about 20,000 s.c.f./b., and the hydrogen consumption rate is from about 6 to about 20 mols of hydrogen per gram atom of sulfur.

13. The process of claim 11 wherein the hydrogenating catalyst is selected from the group consisting of Group VI metals, Group VIII metals, their oxides and sulfides.

14. The process of claim 11 wherein the hydrogenation catalysts consists essentially of the oxides of nickel, cobalt and molybdenum.

15. The process of claim 11 wherein y is less than 0.4.

16. A process for hydrodesulfurizing heavy petroleum hydrocarbons containing asphaltic and metalliferous compounds as contaminants, containing substantial amounts of hydrocarbon components boiling above 300 F., containing residual materials, and containing harmful amounts of sulfur compounds, which process comprises contacting the hydrocarbons with hydrogen at an initial temperature from about 700 to about 800 F., gradually increasing the temperature at a rate sufficient to maintain the sulfur content of the product below a predetermined level and terminating the contacting when a temperature in the range from about 775 to about 875 F. has been reached, while maintaining a pressure from about 300 to about 2000 p.s.i., a space velocity from about 0.1 to about 10.0 volumes of heavy hydrocarbon per volume of catalyst per hour and a hydrogen consumption rate from about 1 to about 5 mols of hydrogen per gram atom of sulfur removed from the heavy hydrocarbons, the contacting being conducted in the presence of a catalyst comprising essentially a minor amount of a hydrogenating catalyst composited with a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than 100 A. radius and having less than 10 percent of said pore volume in pores larger than A. radius prepared by drying and calcining an alumina sol having the formula Al(OI-I) Z wherein Z is the anion of a water soluble aluminum salt, x+y=3 and y is less than 1, said alumina sol being prepared by the reaction of said salt and another source of aluminum ions in the presence of water at a temperature above about 120 F. and a pH below about 4.5.

17. A process for hydrocracking heavy petroleum hydrocarbons containing sulfur, asphaltic and metalliferous compounds as contaminants, containing substantial amounts of hydrocarbon components boiling above 300 F., and containing residual materials, which process comprises ccntacting the hydrocarbons with hydrogen at an initial temperature from about 750 F. to about 850 F., gradually increasing the temperature at a rate sutl'icient to maintain a predetermined rate of conversion and terminating the contacting when a temperature in the range from about 810 F. to about 900 F. has been reached, while maintaining a pressure from about 1500 to about 3000 p.s.i., a space velocity from about 0.1 to about 5.0 volumes of heavy hydrocarbon per volume of catalyst per hour and a hydrogen consumption rate from about 6 to about 40 mols of hydrogen per gram atom of sulfur removed from the heavy hydrocarbons, the contacting being conducted in the presence of a catalyst comprising essentially a minor amount of a hydrogenating catalyst composited with a major amount of an activated alumina having less than 5 percent of its pore volume that is in the form of pores having a radius of 0 to 300 A. in pores larger than A. radius and having less than 10 percent of said pore volume in pores larger than 80 A. radius prepared by drying and calcining an alumina sol having the formula Al(OI-I) Z wherein Z is the anion of a water soluble aluminum salt, x+y=3 and y is less than 1, said alumina sol being prepared by the reaction of said salt and another source of aluminum ions in the presence of water at a temperature above about F. and a pH below about 4.5.

18. The process of claim 16 wherein the initial temperature is from about 750 to about 775 F., the contacting is terminated when a temperature in the range from about 800 to about 850 F. is reached, the pressure -is from about 1000 to about 1500 p.s.i., the space velocity is from about 0.2 to about 3.0 volumes of heavy hydrocarbon per volume of catalyst per hour, the hydrogen consumption rate is from about 2 to about 5 mols of hydrogen per gram atom of sulfur removed from the heavy hydrocarbon, the hydrogenating catalyst is selected from the group consisting of Group VI metals, Group VIII metals, their oxides and sulfides, y is less than 0.4, the water soluble aluminum salt is aluminum chloride, the other source of aluminum ions is aluminum metal and the sol is prepared at a temperature from about F. up to the boiling point of the mixture.

19. The process of claim 17 wherein the initial temperature is from about 780 to about 810 F., the contacting is terminated when a temperature in the range from about 840 to about 875 F. is reached, the pressure is from about 2000 to about 2500 p.s.i., the space velocity is from about 0.2 to about 2.0 volumes of heavy y ar on per volume of catalyst per hour, the hydrogen consumption rate is from about 6 to about 20 mols i 7 l 8 of hydrogen per gram atom of sulfur removed from the References Cited heavy hydrocarbons, the hydrogenating catalyst is selected UNITED STATES PATENTS from the group consisting of Group VI metals, Group VIII metals, their oxides and sulfides, y is less than 0.4, 3,222,273 12/1965 et 208112 the Water soluble aluminum salt is aluminum chloride, 5 the other source of aluminum ions is aluminum metal and DELBERT GANTZ P'mmry Examine the sol is prepared at a temperature from about 160 F. ABRAHAM RIMENS, Examiner. up to the boiling point of the mixture. 

1. A PROCESS FOR HYDROGEN TREATMENT OF HEAVY PETROLEUM HYDROCARBONS CONTAINING SULFUR, ASPHALTIC AND METALLIFEROUS COMPOUNDS AS CONTAMINANTS, CONTAINING SUBSTANTIAL AMOUNTS OF HYDROCARBON COMPONENTS BOILING ABOVE 300*F. AND CONTAINING RESIDUAL MATERIALS, WHICH PROCESS COMPRISES CONTACTING THE HYDROCARBONS WITH HYDROGEN AT A HYDROGEN PARTIAL PRESSURE FROM ABOUT 300 TO ABOUT 3000 P.S.I., A SPACE VELOCITY FROM ABOUT 0.1 TO ABOUT 10.0 VOLUMES OF HEAVY HYDROCARBON PER VOLUME OF CATALYST PER HOUR, A HYDROGEN COMSUMPTION RATE FROM ABOUT 1 TO ABOUT 40 MOLS OF HYDROGEN PER GRAM ATOM OF SULFUR REMOVED FROM THE HEAVY HYDROCARBONS AND AT A TEMPERATURE FROM ABOUT 700* TO ABOUT 900*F. IN THE PRESENCE OF A CATALYST COMPRISING ESSENTIALLY A MINOR AMOUNT OF A HYDROGENATING CATALYST COMPOSITED WITH A MAJOR AMOUNT OF AN ACTIVATED ALUMINA HAVING LESS THAN 5 PERCENT OF ITS PORE VOLUME THAT IS IN THE FORM OF PORES HAVING A RADIUS OF 0 TO 300 A. IN PORES LARGER THAN 100 A. RADIUS AND HAVING LESS THAN 10 PERCENT OF SAID PORE VOLUME IN PORES LARGER THAN 80 A. RADIUS PREPARED BY DRYING AND CALCINING AN ALUMINUM SOL HAVING THE FORMULA AL(OH)XZY, WHEREIN Z IS THE ANION OF WATER SOLUBLE ALUMINUM SALT, X+Y=3 AND Y IS LESS THAN 1, SAID ALUMINA SOL BEING PREPARED BY THE REACTION OF SAID SALT AND ANOTHER SOURCE OF ALUMINUM IONS IN THE PRESENCE OF WATER AT A TEMPERATURE ABOVE ABOUT 120*F. AND AT A PH BELOW ABOUT 4.5. 